Process for converting gaseous alkanes to liquid hydrocarbons

ABSTRACT

Embodiments disclose a process for converting gaseous alkanes to higher molecular weight hydrocarbons, olefins or mixtures thereofs wherein a gaseous feed containing alkanes may be reacted with a dry bromine vapor to form alkyl bromides and hydrobromic acid vapor. The mixture of alkyl bromides and hydrobromic acid then may be reacted over a synthetic crystalline alumino-silicate catalyst, such as a ZSM-5 or an X or Y type zeolite, at a temperature of from about 250° C. to about 500° C. so as to form hydrobromic acid vapor and higher molecular weight hydrocarbons, olefins or mixtures thereof. Various methods are disclosed to remove the hydrobromic acid vapor from the higher molecular weight hydrocarbons, olefins or mixtures thereof and to generate bromine from the hydrobromic acid for use in the process.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates to a process for converting lowermolecular weight, gaseous alkanes to olefins, higher molecular weighthydrocarbons, or mixtures thereof that may be useful as fuels ormonomers and intermediaries in the production of fuels or chemicals,such as lubricant and fuel additives, and more particularly, in one ormore embodiments, to a process wherein a gas containing lower molecularweight alkanes is reacted with a dry bromine vapor to form alkylbromides and hydrobromic acid which in turn are reacted over acrystalline alumino-silicate catalyst to form olefins, higher molecularweight hydrocarbons or mixtures thereof.

2. Description of Related Art

Natural gas, which is primarily composed of methane and other lightalkanes, has been discovered in large quantities throughout the world.Many of the locales in which natural gas has been discovered are farfrom populated regions which have significant gas pipelineinfrastructure or market demand for natural gas. Due to the low densityof natural gas, transportation thereof in gaseous form by pipeline or ascompressed gas in vessels is expensive. Accordingly, practical andeconomic limits exist to the distance over which natural gas may betransported in gaseous form. Cryogenic liquefaction of natural gas (LNG)is often used to more economically transport natural gas over largedistances. However, this LNG process is expensive and there are limitedregasification facilities in only a few countries that are equipped toimport LNG

Another use of methane is as feed to processes for the production ofmethanol. Methanol is made commercially via conversion of methane tosynthesis gas (CO and H₂) at high temperatures (approximately 1000° C.)followed by synthesis at high pressures (approximately 100 atmospheres).There are several types of technologies for the production of synthesisgas from methane. Among these are steam-methane reforming (SMR), partialoxidation (POX), autothermal reforming (ATR), gas-heated reforming(GHR), and various combinations thereof. SMR and GHR operate at highpressures and temperatures, generally in excess of 600° C., and requireexpensive furnaces or reactors containing special heat andcorrosion-resistant alloy tubes filled with expensive reformingcatalyst. POX and ATR processes operate at high pressures and evenhigher temperatures, generally in excess of 1000° C. As there are noknown practical metals or alloys that can operate at these temperatures,complex and costly refractory-lined reactors and high-pressurewaste-heat boilers to quench and cool the synthesis gas effluent arerequired. Also, significant capital cost and large amounts of power arerequired for compression of oxygen or air to these high-pressureprocesses. Thus, due to the high temperatures and pressures involved,synthesis gas technology is expensive, resulting in a high cost methanolproduct which limits higher-value uses thereof, such as for chemicalfeedstocks and solvents. Furthermore production of synthesis gas isthermodynamically and chemically inefficient, producing large excessesof waste heat and unwanted carbon dioxide, which tends to lower theconversion efficiency of the overall process. Fischer-TropschGas-to-Liquids (GTL) technology can also be used to convert synthesisgas to heavier liquid hydrocarbons, however investment cost for thisprocess is even higher. In each case, the production of synthesis gasrepresents a large fraction of the capital costs for these methaneconversion processes.

Numerous alternatives to the conventional production of synthesis gas asa route to methanol or synthetic liquid hydrocarbons have been proposed.However, to date, none of these alternatives has attained commercialstatus for various reasons. Some of the previous alternative prior-artmethods, such as disclosed in U.S. Pat. No. 5,243,098 or 5,334,777 toMiller, teach reacting a lower alkane, such as methane, with a metallichalide to form a metal halide and hydrohalic acid which are in turnreduced with magnesium oxide to form the corresponding alkanol. However,halogenation of methane using chlorine as the preferred halogen resultsin poor selectivity to the monomethyl halide (CH₃Cl), resulting inunwanted by-products such as CH₂Cl₂ and CHCl₃ which are difficult toconvert or require severe limitation of conversion per pass and hencevery high recycle rates.

Other prior art processes propose the catalytic chlorination orbromination of methane as an alternative to generation of synthesis gas(CO and H₂). To improve the selectivity of a methane halogenation stepin an overall process for the production of methanol, U.S. Pat. No.5,998,679 to Miller teaches the use of bromine, generated by thermaldecomposition of a metal bromide, to brominate alkanes in the presenceof excess alkanes, which results in improved selectivity tomono-halogenated intermediates such as methyl bromide. To avoid thedrawbacks of utilizing fluidized beds of moving solids, the processutilizes a circulating liquid mixture of metal chloride hydrates andmetal bromides. Processes described in U.S. Pat. No. 6,462,243 B1, U.S.Pat. No. 6,472,572 B1, and U.S. Pat. No. 6,525,230 to Grosso are alsocapable of attaining higher selectivity to mono-halogenatedintermediates by the use of bromination. The resulting alkyl bromideintermediates such as methyl bromide, are further converted to thecorresponding alcohols and ethers, by reaction with metal oxides incirculating beds of moving solids. Another embodiment of U.S. Pat. No.6,525,230 avoids the drawbacks of moving beds by utilizing a zonedreactor vessel containing a fixed bed of metal bromide/oxide solids thatis operated cyclically in four steps. While certain ethers, such asdimethyl ether (“DME”) are a promising potential diesel engine fuelsubstitute, as of yet, there currently exists no substantial market forDME, and hence an expensive additional catalytic process conversion stepwould be required to convert DME into a currently marketable product.Other processes have been proposed which circumvent the need forproduction of synthesis gas, such as U.S. Pat. No. 4,467,130 to Olah inwhich methane is catalytically condensed into gasoline-rangehydrocarbons via catalytic condensation using superacid catalysts.However, none of these earlier alternative approaches have resulted incommercial processes.

It is known that substituted alkanes, in particular methanol, can beconverted to olefins and gasoline boiling-range hydrocarbons overvarious forms of crystalline alumino-silicates also known as zeolites.In the Methanol to Gasoline (MTG) process, a shape selective zeolitecatalyst, ZSM-5, is used to convert methanol to gasoline. Coal ormethane gas can thus be converted to methanol using conventionaltechnology and subsequently converted to gasoline. However due to thehigh cost of methanol production, and at current or projected prices forgasoline, the MTG process is not considered economically viable. Thus, aneed exists for an economic process for the conversion of methane andother alkanes found in natural gas to olefins, higher molecular weighthydrocarbons or mixtures thereof which, due to their higher density andvalue, are more economically transported thereby significantly aidingdevelopment of remote natural gas reserves. Further, a need exists forsuch a process that is relatively inexpensive, safe and simple.

SUMMARY OF THE INVENTION

To achieve the foregoing and other objects, and in accordance with thepurposes of the present invention, as embodied and broadly describedherein, one characterization of the present invention is a processcomprising: separating hydrobromic acid from a gaseous stream comprisinghydrobromic acid and hydrocarbons; converting said hydrobromic acid toat least bromine; and contacting said bromine with gaseous alkanes toform bromination products comprising alkyl bromides.

In another characterization of the present invention, a process isprovided that comprises: contacting a gaseous stream comprisinghydrobromic acid and hydrocarbons with an aqueous solution comprising abase selected from the group consisting of a metal hydroxide, a metaloxy-bromide species, and combinations thereof such that the hydrobromicacid is neutralized to form a metal bromide salt in the aqueoussolution; oxidizing said aqueous solution containing said metal bromidesalt to form oxidation products comprising bromine and said base;separating said bromine from said aqueous solution comprising said base;and contacting said bromine with gaseous alkanes to form alkyl bromides.

In another characterization of the present invention, a process isprovided that comprises: contacting a gaseous stream comprisinghydrobromic acid and hydrocarbons with water, wherein said hydrobromicacid dissolves in said water to form an aqueous solution comprising saidwater and said hydrobromic acid; neutralizing said hydrobromic acid toform a metal bromide salt; oxidizing said metal bromide salt to form anoxidation product comprising bromine; and contacting said bromine withgaseous alkanes to form bromination products comprising alkyl bromides.

In another characterization of the present invention, a process isprovided that comprises: reacting hydrobromic acid with a metal oxide toform reaction products comprising a metal bromide and steam, whereinsaid hydrobromic acid is contained in a gaseous stream comprising saidhydrobromic acid and hydrocarbons; reacting said metal bromide with agas comprising oxygen to form reaction products comprising bromine andsaid metal oxide; and contacting said bromine with gaseous alkanes toform bromination products comprising alkyl bromides.

BRIEF DESCRIPTION OF THE DRAWINGS

The accompanying drawings, which are incorporated in and form a part ofthe specification, illustrate the embodiments of the present inventionand, together with the description, serve to explain the principles ofthe invention.

In the drawings:

FIG. 1 is a simplified block flow diagram of an embodiment of theprocess of the present invention;

FIG. 2 is a schematic view of one embodiment of the process of thepresent invention;

FIG. 3 is a schematic view of another embodiment of process of thepresent invention;

FIG. 4A is schematic view of another embodiment of the process of thepresent invention;

FIG. 4B is a schematic view of the embodiment of the process of thepresent invention illustrated in FIG. 4A depicting an alternativeprocessing scheme which may be employed when oxygen is used in lieu ofair in the oxidation stage;

FIG. 5A is a schematic view of the embodiment of the process of thepresent invention illustrated in FIG. 4A with the flow through the metaloxide beds being reversed;

FIG. 5B is a schematic view of the embodiment of the process of thepresent invention illustrated in FIG. 5A depicting an alternativeprocessing scheme which may be employed when oxygen is used in lieu ofair in the oxidation stage;

FIG. 6A is a schematic view of another embodiment of the process of thepresent invention;

FIG. 6B is a schematic view of the embodiment of the process of thepresent invention illustrated in FIG. 6A depicting an alternativeprocessing scheme which may be employed when oxygen is used in lieu ofair in the oxidation stage;

FIG. 7 is a schematic view of another embodiment of the process of thepresent invention;

FIG. 8 is a schematic view of the embodiment of the process of thepresent invention illustrated in FIG. 7 with the flow through the metaloxide beds being reversed; and

FIG. 9 is a schematic view of another embodiment of the process of thepresent invention.

FIG. 10 is a graph of methyl bromide conversion and product selectivityfor the oligomerization reaction of the process of the present inventionas a function of temperature;

FIG. 11 is a graph comparing conversion and selectivity for the exampleof methyl bromide, dry hydrobromic acid and methane versus only methylbromide plus methane;

FIG. 12 is a graph of product selectivity from reaction of methylbromide and dibromomethane vs. product selectivity from reaction ofmethyl bromide only;

FIG. 13 is a graph of a Paraffinic Olefinic Napthenic and Aromatic(PONA) analysis of a typical condensed product sample of the process ofthe present invention; and

FIG. 14 is a graph of a PONA analysis of another typical condensedproduct sample of the present invention.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

As utilized throughout this description, the term “lower molecularweight alkanes” refers to methane, ethane, propane, butane, pentane ormixtures thereof. As also utilized throughout this description, “alkylbromides” refers to mono, di, and tri brominated alkanes. Also, the feedgas in lines 11 and 111 in the embodiments of the process of the presentinvention as illustrated in FIGS. 2 and 3, respectively, is preferablynatural gas which may be treated to remove sulfur compounds and carbondioxide. In any event, it is important to note that small amounts ofcarbon dioxide, e.g. less than about 2 mol %, can be tolerated in thefeed gas to the process of the present invention.

A block flow diagram generally depicting an embodiment of a process ofthe present invention is illustrated in FIG. 1, while specificembodiments of the process illustrated in FIG. 1 are illustrated inFIGS. 2 and 3. Referring to FIG. 1, a gas stream comprising recycle gasand a natural feed gas is combined with dry bromine vapor and fed to analkane bromination reactor. The recycle gas and the natural gas feed maycomprise lower molecular weight hydrocarbons. In the alkane brominationreactor, the gas stream and the dry bromine vapor are reacted to producegaseous alkyl bromides and hydrobromic acid vapors. As illustrated,gaseous alkyl bromides and hydrobromic acid vapors are fed to the alkylbromide conversion reactor. In the alkyl bromide conversion reactor, thegaseous alkyl bromides are reacted to form higher molecular weighthydrocarbons and additional hydrobromic acid vapors. In the illustratedembodiment, the hydrobromic acid vapors are then removed from the highermolecular hydrocarbons in the hydrobromic acid removal unit by arecirculated aqueous solution. As illustrated in FIG. 1, therecirculated aqueous solution carries the hydrobromic acid (or metalbromide salt if the acid is neutralized by the aqueous solution) to thebromide oxidation unit. As will be discussed in more detail below, thehydrobromic acid may be neutralized in the bromide oxidation unit toform a metal bromide salt. Oxygen or air is supplied to the bromideoxidation unit to oxidize the metal bromide salt to form the bromine,which is then recycled to the alkane bromination reactor.

In the illustrated embodiment, a natural gas feed is also introducedinto the hydrobromic acid removal unit. From the hydrobromic acidremoval unit, the natural gas feed and the higher molecular hydrocarbonsare fed to the dehydration and product recovery unit. In the dehydrationand product recovery unit, water is removed from the higher molecularweight hydrocarbons and a hydrocarbon liquid product is produced. Inaddition, a gas stream of recycle gas and the natural gas feed areconveyed to the alkane bromination reactor. Accordingly, the processillustrated in FIG. 1 may be used to produce a liquid hydrocarbonproduct from lower molecular hydrocarbons.

Referring to FIG. 2, a gas stream containing lower molecular weightalkanes, comprised of a mixture of a feed gas plus a recycled gas streamat a pressure in the range of about 1 bar to about 30 bar, istransported or conveyed via line, pipe or conduit 62, mixed with drybromine liquid transported via line 25 and pump 24, and passed to heatexchanger 26 wherein the liquid bromine is vaporized. The mixture oflower molecular weight alkanes and dry bromine vapor is fed to reactor30. Preferably, the molar ratio of lower molecular weight alkanes to drybromine vapor in the mixture introduced into reactor 30 is in excess of2.5:1. Reactor 30 has an inlet pre-heater zone 28 which heats themixture to a reaction initiation temperature in the range of about 250°C. to about 400° C.

In first reactor 30, the lower molecular weight alkanes are reactedexothermically with dry bromine vapor at a relatively low temperature inthe range of about 250° C. to about 600° C., and at a pressure in therange of about 1 bar to about 30 bar to produce gaseous alkyl bromidesand hydrobromic acid vapors. The upper limit of the operatingtemperature range is greater than the upper limit of the reactioninitiation temperature range to which the feed mixture is heated due tothe exothermic nature of the bromination reaction. In the case ofmethane, the formation of methyl bromide occurs in accordance with thefollowing general reaction:CH₄(g)+Br₂(g)→CH₃Br(g)+HBr(g)

This reaction occurs with a significantly high degree of selectivity tomethyl bromide. For example, in the case of bromination of methane, amethane to bromine ratio of about 4.5:1 increases the selectivity to themono-halogenated methyl bromide to that obtained using smaller methaneto bromine ratios. Small amounts of dibromomethane and tribromomethaneare also formed in the bromination reaction. Higher alkanes, such asethane, propane and butane, are also readily brominated resulting inmono and multiple brominated species such as ethyl bromides, propylbromides and butyl bromides. If an alkane to bromine ratio ofsignificantly less than about 2.5 to 1 is utilized, a lower selectivityto methyl bromide occurs and significant formation of undesirable carbonsoot is observed. Further, the dry bromine vapor that is feed into firstreactor 30 is substantially water-free. Applicant has discovered thatelimination of substantially all water vapor from the bromination stepin first reactor 30 substantially eliminates the formation of unwantedcarbon dioxide thereby increasing the selectivity of alkane brominationto alkyl bromides and eliminating the large amount of waste heatgenerated in the formation of carbon dioxide from alkanes.

The effluent that contains alkyl bromides and hydrobromic acid iswithdrawn from the first reactor via line 31 and is partially cooled inheat exchanger 32 before flowing to a second reactor 34. The temperatureto which the effluent is partially cooled in heat exchanger 34 is in therange of about 150° C. to about 350° C. when it is desired to convertthe alkyl bromides to higher molecular weight hydrocarbons in secondreactor 34, or to range of about 150° C. to about 450° C. when it isdesired to convert the alkyl bromides to olefins a second reactor 34. Insecond reactor 34, the alkyl bromides are reacted exothermically over afixed bed 33 of crystalline alumino-silicate catalyst, preferably azeolite catalyst. The temperature and pressure employed in secondreactor, as well as the zeolite catalyst, will determine the product(s)that is formed from the reaction of alkyl bromides occurring in secondreactor 34.

The crystalline alumino-silicate catalyst employed in second reactor 34is preferably a zeolite catalyst and most preferably a ZSM-5 zeolitecatalyst when it is desired to form higher molecular weighthydrocarbons, Although the zeolite catalyst is preferably used in thehydrogen, sodium or magnesium form, the zeolite may also be modified byion exchange with other alkali metal cations, such as Li, Na, K or Cs,with alkali-earth metal cations, such as Mg, Ca, Sr or Ba, or withtransition metal cations, such as Ni, Mn, V, W, or to the hydrogen form.Other zeolite catalysts having varying pore sizes and acidities, whichare synthesized by varying the alumina-to-silica ratio may be used inthe second reactor 34 as will be evident to a skilled artisan.

When it is desired to form olefins from the reaction of alkyl bromidesin reactor 34, the crystalline alumino-silicate catalyst employed insecond reactor 34 is preferably a zeolite catalyst, and most preferablyan X type or Y type zeolite catalyst. A preferred zeolite is 10 X or Ytype zeolite, although other zeolites with differing pore sizes andacidities, which are synthesized by varying the alumina-to-silica ratiomay be used in the process of the present invention as will be evidentto a skilled artisan. Although the zeolite catalyst is preferably usedin a protonic form, a sodium form or a mixed protonic/sodium form, thezeolite may also be modified by ion exchange with other alkali metalcations, such as Li, K or Cs, with alkali-earth metal cations, such asMg, Ca, Sr or Ba, or with transition metal cations, such as Ni, Mn, V,W, or to the hydrogen form. These various alternative cations have aneffect of shifting reaction selectivity. Other zeolite catalysts havingvarying pore sizes and acidities, which are synthesized by varying thealumina-to-silica ratio may be used in the second reactor 34 as will beevident to a skilled artisan.

The temperature at which the second reactor 34 is operated is animportant parameter in determining the selectivity of the reaction tohigher molecular hydrocarbons or to olefins.

Where a catalyst is selected to form higher molecular weighthydrocarbons in reactor 34, it is preferred to operate second reactor 34at a temperature within the range of about 150° to 450°. Temperaturesabove about 300° C. in the second reactor result in increased yields oflight hydrocarbons, such as undesirable methane, whereas lowertemperatures increase yields of heavier molecular weight hydrocarbonproducts. At the low end of the temperature range, with methyl bromidereacting over ZSM-5 zeolite at temperatures as low as 150° C.significant methyl bromide conversion on the order of 20% is noted, witha high selectivity towards C₅+ products. Notably, in the case of thealkyl bromide reaction over the preferred zeolite ZSM-5 catalyst,cyclization reactions also occur such that the C7+ fractions arecomposed primarily of substituted aromatics. At increasing temperaturesapproaching 300° C., methyl bromide conversion increases towards 90% orgreater, however selectivity towards C₅+ products decreases andselectivity towards lighter products, particularly undesirable methane,increases. Surprisingly, very little ethane or C₂,-C₃ olefin componentsare formed. At temperatures approaching 450° C., almost completeconversion of methyl bromide to methane occurs. In the optimum operatingtemperature range of between about 300° C. and 400° C., as a byproductof the reaction, a small amount of carbon will build up on the catalystover time during operation, causing a decline in catalyst activity overa range of hours, up to hundreds of hours, depending on the reactionconditions and the composition of the feed gas. It is believed thathigher reaction temperatures above about 400° C., associated with theformation of methane favor the thermal cracking of alkyl bromides andformation of carbon or coke and hence an increase in the rate ofdeactivation of the catalyst. Conversely, temperatures at the lower endof the range, particularly below about 300° C. may also contribute tocoking due to a reduced rate of desorption of heavier products from thecatalyst. Hence, operating temperatures within the range of about 150°C. to about 450° C., but preferably in the range of about 300° C. toabout 400° C. in the second reactor 34 balance increased selectivity ofthe desired C₅+ products and lower rates of deactivation due to carbonformation, against higher conversion per pass, which minimizes thequantity of catalyst, recycle rates and equipment size required.

Where a catalyst is selected to form olefins in reactor 34, it ispreferred to operate second reactor 34 at a temperature within the rangeof about 250° C. to 500° C. Temperatures above about 450° C. in thesecond reactor can result in increased yields of light hydrocarbons,such as undesirable methane and also deposition of coke, whereas lowertemperatures increase yields of ethylene, propylene, butylene andheavier molecular weight hydrocarbon products. Notably, in the case ofthe alkyl bromide reaction over the preferred 10 X zeolite catalyst, itis believed that cyclization reactions also occur such that the C7+fractions contain substantial substituted aromatics. At increasingtemperatures approaching 400° C., it is believed that methyl bromideconversion increases towards 90% or greater, however selectivity towardsC₅+ products decreases and selectivity towards lighter products,particularly olefins increases. At temperatures exceeding 550° C., it isbelieved that a high conversion of methyl bromide to methane andcarbonaceous, coke occurs. In the preferred operating temperature rangeof between about 300° C. and 450° C., as a byproduct of the reaction, alesser amount of coke probably will build up on the catalyst over timeduring operation, causing a decline in catalyst activity over a range ofhours, up to hundreds of hours, depending on the reaction conditions andthe composition of the feed gas. It is believed that higher reactiontemperatures above about 400° C., associated with the formation ofmethane favor the thermal cracking of alkyl bromides and formation ofcarbon or coke and hence an increase in the rate of deactivation of thecatalyst. Conversely, temperatures at the lower end of the range,particularly below about 300° C. may also contribute to coking due to areduced rate of desorption of heavier products from the catalyst. Hence,operating temperatures within the range of about 250° C. to about 500°C., but preferably in the range of about 300° C. to about 450° C. in thesecond reactor 34 balance increased selectivity of the desired olefinsand C₅+ products and lower rates of deactivation due to carbonformation, against higher conversion per pass, which minimizes thequantity of catalyst, recycle rates and equipment size required.

The catalyst may be periodically regenerated in situ, by isolatingreactor 34 from the normal process flow, purging with an inert gas vialine 70 at a pressure in a range from about 1 to about 5 bar at anelevated temperature in the range of about 400° C. to about 650° C. toremove unreacted material adsorbed on the catalyst insofar as ispractical, and then subsequently oxidizing the deposited carbon to CO₂by addition of air or inert gas-diluted oxygen to reactor 34 via line 70at a pressure in the range of about 1 bar to about 5 bar at an elevatedtemperature in the range of about 400° C. to about 650° C. Carbondioxide and residual air or inert gas is vented from reactor 34 via line75 during the regeneration period.

The effluent which comprises hydrobromic acid and higher molecularweight hydrocarbons, olefins or mixtures thereof is withdrawn from thesecond reactor 34 via line 35 and is cooled to a temperature in therange of 0° C. to about 100° C. in exchanger 36 and combined with vaporeffluent in line 12 from hydrocarbon stripper 47, which contains feedgas and residual higher molecular weight hydrocarbons stripped-out bycontact with the feed gas in hydrocarbon stripper 47. The combined vapormixture is passed to a scrubber 38 and contacted with a concentratedaqueous partially-oxidized metal bromide salt solution containing metalhydroxide, metal oxide, metal oxy-bromide or mixtures of these species,which is transported to scrubber 38 via line 41. The preferred metal ofthe bromide salt is Fe(III), Cu(II) or Zn(II), or mixtures thereof, asthese are less expensive and readily oxidize at lower temperatures inthe range of about 120° C. to about 180° C., allowing the use ofglass-lined or fluorpolymer-lined equipment; although Co(II), Ni(II),Mn(II), V(II), Cr(II) or other transition-metals which form oxidizablebromide salts may be used in the process of the present invention.Alternatively, alkaline-earth metals which also form oxidizable bromidesalts, such as Ca(II) or Mg(II) may be used. Any liquid hydrocarbonscondensed in scrubber 38 may be skimmed and withdrawn in line 37 andadded to liquid hydrocarbons exiting the product recovery unit 52 inline 54. Hydrobromic acid is dissolved in the aqueous solution andneutralized by the metal hydroxide, metal oxide, metal oxy-bromide ormixtures of these species to yield metal bromide salt in solution andwater which is removed from the scrubber 38 via line 44.

The residual vapor phase containing olefins, higher molecular weighthydrocarbons or mixtures thereof that is removed as effluent from thescrubber 38 is forwarded via line 39 to dehydrator 50 to removesubstantially all water via line 53 from the vapor stream. The water isthen removed from the dehydrator 50 via line 53. The dried vapor streamcontaining olefins, higher molecular weight hydrocarbons or mixturesthereof is further passed via line 51 to product recovery unit 52 torecover olefins, the C₅+ gasoline-range hydrocarbon fraction or mixturesthereof as a liquid product in line 54. Any conventional method ofdehydration and liquids recovery, such as solid-bed desiccant adsorptionfollowed by refrigerated condensation, cryogenic expansion, orcirculating absorption oil or other solvent, as used to process naturalgas or refinery gas streams, and/or to recover olefinic hydrocarbons, aswill be evident to a skilled artisan, may be employed in the process ofthe present invention. The residual vapor effluent from product recoveryunit 52 is then split into a purge stream 57 which may be utilized asfuel for the process and a recycled residual vapor which is compressedvia compressor 58. The recycled residual vapor discharged fromcompressor 58 is split into two fractions. A first fraction that isequal to at least 2.5 times the feed gas molar volume is transported vialine 62 and is combined with dry liquid bromine conveyed by pump 24,heated in exchanger 26 to vaporize the bromine and fed into firstreactor 30. The second fraction is drawn off of line 62 via line 63 andis regulated by control valve 60, at a rate sufficient to dilute thealkyl bromide concentration to reactor 34 and absorb the heat ofreaction such that reactor 34 is maintained at the selected operatingtemperature, preferably in the range of about 300° C. to about 450° C.in order to maximize conversion versus selectivity and to minimize therate of catalyst deactivation due to the deposition of carbon. Thus, thedilution provided by the recycled vapor effluent permits selectivity ofbromination in the first reactor 30 to be controlled in addition tomoderating the temperature in second reactor 34.

Water containing metal bromide salt in solution which is removed fromscrubber 38 via line 44 is passed to hydrocarbon stripper 47 whereinresidual dissolved hydrocarbons are stripped from the aqueous phase bycontact with incoming feed gas transported via line 11. The strippedaqueous solution is transported from hydrocarbon stripper 47 via line 65and is cooled to a temperature in the range of about 0° C. to about 70°C. in heat exchanger 46 and then passed to absorber 48 in which residualbromine is recovered from vent stream in line 67. The aqueous solutioneffluent from adsorber 48 is transported via line 49 to a heat exchanger40 to be preheated to a temperature in the range of about 100° C. toabout 600° C., and most preferably in the range of about 120° C. toabout 180° C. and passed to third reactor 16. Oxygen or air is deliveredvia line 10 by blower or compressor 13 at a pressure in the range ofabout ambient to about 5 bar to bromine stripper 14 to strip residualbromine from water. Water is removed from stripper 14 in line 64 andcombined with water stream 53 from dehydrator 50 to form water effluentstream in line 56 which is removed from the process. The oxygen or airleaving bromine stripper 14 is fed via line 15 to reactor 16 whichoperates at a pressure in the range of about ambient to about 5 bar andat a temperature in the range of about 100° C. to about 600° C., butmost preferably in the range of about 120° C. to about 180° C. so as tooxidize an aqueous metal bromide salt solution to yield elementalbromine and metal hydroxide, metal oxide, metal oxy-bromide or mixturesof these species. As stated above, although Co(II), Ni(II), Mn(II),V(II), Cr(II) or other transition-metals which form oxidizable bromidesalts can be used, the preferred metal of the bromide salt is Fe(III),Cu(II), or Zn(II), or mixtures thereof, as these are less expensive andreadily oxidize at lower temperatures in the range of about 120° C. toabout 180° C., which should allow the use of glass-lined orfluorpolymer-lined equipment. Alternatively, alkaline-earth metals whichalso form oxidizable bromide salts, such as Ca(II) or Mg(II) could beused.

Hydrobromic acid reacts with the metal hydroxide, metal oxide, metaloxy-bromide or mixtures of these species so formed to once again yieldthe metal bromide salt and water. Heat exchanger 18 in reactor 16supplies heat to vaporize water and bromine. Thus, it is believed thatthe overall reactions result in the net oxidation of hydrobromic acidproduced in first reactor 30 and second reactor 34 to elemental bromineand steam in the liquid phase catalyzed by the metal bromide/metal oxideor metal hydroxide operating in a catalytic cycle. In the case of themetal bromide being Fe(III)Br3, the reactions are believed to be:Fe(+3a)+6Br(−a)+3H(+a)+3/2O₂(g)=3Br₂(g)+Fe(OH)₃  1)3HBr(g)+H₂O=3H(+a)+3Br(−a)+H₂O  2)3H(+a)+3Br(−a)+Fe(OH)₃=Fe(+3a)+3Br(−a)+3H₂O  3)In the case of the metal bromide being CU(II)Br2, the reactions arebelieved to be:4Cu(+2a)+8Br(−a)+3H₂0+3/2O₂(g)=3Br₂(g)+CuBr₂.3Cu(OH)₂  1)6HBr(g)+H₂O=6H(+a)+6Br(−a)+H₂O  2)6H(+a)+6Br(−a)+CuBr₂.3Cu(OH)₂=4Cu(+2a)+8Br(−a)+6H₂O  3)

The elemental bromine and water and any residual oxygen or nitrogen (ifair is utilized as the oxidant) leaving as vapor from the outlet ofthird reactor 16 via line 19, are cooled in condenser 20 at atemperature in the range of about 0° C. to about 70° C. and a pressurein the range of about ambient to 5 bar to condense the bromine and waterand passed to three-phase separator 22. In three-phase separator 22,since liquid water has a limited solubility for bromine, on the order ofabout 3% by weight, any additional bromine which is condensed forms aseparate, denser liquid bromine phase. The liquid bromine phase,however, has a notably lower solubility for water, on the order of lessthan 0.1%. Thus a substantially dry bromine vapor can be easily obtainedby condensing liquid bromine and water, decanting water by simplephysical separation and subsequently re-vaporizing liquid bromine.

Liquid bromine is pumped in line 25 from three-phase separator 22 viapump 24 to a pressure sufficient to mix with vapor stream 62. Thusbromine is recovered and recycled within the process. The residualoxygen or nitrogen and any residual bromine vapor which is not condensedexits three-phase separator 22 and is passed via line 23 to brominescrubber 48, wherein residual bromine is recovered by solution into andby reaction with reduced metal bromides in the aqueous metal bromidesolution stream 65. Water is removed from separator 22 via line 27 andintroduced into stripper 14.

In another embodiment of the invention, referring to FIG. 3, a gasstream containing lower molecular weight alkanes, comprised of mixtureof a feed gas plus a recycled gas stream at a pressure in the range ofabout 1 bar to about 30 bar, is transported or conveyed via line, pipeor conduit 162, mixed with dry bromine liquid transported via pump 124and passed to heat exchanger 126 wherein the liquid bromine isvaporized. The mixture of lower molecular weight alkanes and dry brominevapor is fed to reactor 130. Preferably, the molar ratio of lowermolecular weight alkanes to dry bromine vapor in the mixture introducedinto reactor 130 is in excess of 2.5:1. Reactor 130 has an inletpre-heater zone 128 which heats the mixture to a reaction initiationtemperature in the range of about 250° C. to about 400° C. In firstreactor 130, the lower molecular weight alkanes are reactedexothermically with dry bromine vapor at a relatively low temperature inthe range of about 250° C. to about 600° C., and at a pressure in therange of about 1 bar to about 30 bar to produce gaseous alkyl bromidesand hydrobromic acid vapors. The upper limit of the operatingtemperature range is greater than the upper limit of the reactioninitiation temperature range to which the feed mixture is heated due tothe exothermic nature of the bromination reaction. In the case ofmethane, the formation of methyl bromide occurs in accordance with thefollowing general reaction:CH₄(g)+Br₂(g)→CH₃Br(g)+HBr(g)

This reaction occurs with a significantly high degree of selectivity tomethyl bromide. For example, in the case of bromination of methane, amethane to bromine ratio of about 4.5:1 increases the selectivity to themono-halogenated methyl bromide. Small amounts of dibromomethane andtribromomethane are also formed in the bromination reaction. Higheralkanes, such as ethane, propane and butane, are also readily brominatedresulting in mono and multiple brominated species such as ethylbromides, propyl bromides and butyl bromides. If an alkane to bromineratio of significantly less than about 2.5 to 1 is utilized, a lowerselectivity to methyl bromide occurs and significant formation ofundesirable carbon soot is observed. Further, the dry bromine vapor thatis feed into first reactor 30 is preferably substantially water-free.Applicant has discovered that elimination of substantially all watervapor from the bromination step in first reactor 30 substantiallyeliminates the formation of unwanted carbon dioxide thereby increasingthe selectivity of alkane bromination to alkyl bromides and eliminatingthe large amount of waste heat generated in the formation of carbondioxide from alkanes.

The effluent that contains alkyl bromides and hydrobromic acid iswithdrawn from the first reactor via line 131 and is partially cooled inheat exchanger 132 before flowing to a second reactor 134. Thetemperature to which the effluent is partially cooled in heat exchanger134 is in the range of about 150° C. to about 350° C. where it isdesired to convert the alkyl bromides to higher molecular weighthydrocarbons in second reactor 134, or to range of about 150° C. toabout 450° C. where it is desired to convert the alkyl bromides toolefins in second reactor 134. In second reactor 134, the alkyl bromidesare reacted exothermically over a fixed bed 133 of crystallinealumino-silicate catalyst, preferably a zeolite catalyst. Thetemperature and pressure employed in second reactor 134, as well as thezeolite catalyst, will determine the product that is formed from thereaction of alkyl bromides occurring in second reactor 134.

The crystalline alumino-silicate catalyst employed in second reactor 134is preferably a zeolite catalyst and most preferably a ZSM-5 zeolitecatalyst when it is desired to form higher molecular weighthydrocarbons, Although the zeolite catalyst is preferably used in thehydrogen, sodium or magnesium form, the zeolite may also be modified byion exchange with other alkali metal cations, such as Li, Na, K or Cs,with alkali-earth metal cations, such as Mg, Ca, Sr or Ba, or withtransition metal cations, such as Ni, Mn, V, W, or to the hydrogen form.Other zeolite catalysts having varying pore sizes and acidities, whichare synthesized by varying the alumina-to-silica ratio may be used inthe second reactor 134 as will be evident to a skilled artisan.

When it is desired to form olefins from the reaction of alkyl bromidesin reactor 134, the crystalline alumino-silicate catalyst employed insecond reactor 134 is preferably a zeolite catalyst and most preferablyan X type or Y type zeolite catalyst. A preferred zeolite is 10 X or Ytype zeolite, although other zeolites with differing pore sizes andacidities, which are synthesized by varying the alumina-to-silica ratiomay be used in the process of the present invention as will be evidentto a skilled artisan. Although the zeolite catalyst is preferably usedin a protonic form, a sodium form or a mixed protonic/sodium form, thezeolite may also be modified by ion exchange with other alkali metalcations, such as Li, K or Cs, with alkali-earth metal cations, such asMg, Ca, Sr or Ba, or with transition metal cations, such as Ni, Mn, V,W, or to the hydrogen form. These various alternative cations have aneffect of shifting reaction selectivity. Other zeolite catalysts havingvarying pore sizes and acidities, which are synthesized by varying thealumina-to-silica ratio may be used in the second reactor 134 as will beevident to a skilled artisan.

The temperature at which the second reactor 134 is operated is animportant parameter in determining the selectivity of the reaction tohigher molecular weight hydrocarbons or to olefins.

When a catalyst is selected to form higher molecular weight hydrocarbonsin reactor 134, it is preferred to operate second reactor 134 at atemperature within the range of about 150° to 450°. Temperatures aboveabout 300° C. in the second reactor result in increased yields of lighthydrocarbons, such as undesirable methane, whereas lower temperaturesincrease yields of heavier molecular weight hydrocarbon products. At thelow end of the temperature range, with methyl bromide reacting overZSM-5 zeolite at temperatures as low as 150° C. significant methylbromide conversion on the order of 20% is noted, with a high selectivitytowards C₅+ products. Notably, in the case of the alkyl bromide reactionover the preferred zeolite ZSM-5 catalyst, cyclization reactions alsooccur such that the C7+ fractions are composed primarily of substitutedaromatics. At increasing temperatures approaching 300° C., methylbromide conversion increases towards 90% or greater, however selectivitytowards C₅+ products decreases and selectivity towards lighter products,particularly undesirable methane, increases. Surprisingly, very littleethane or C₂,-C₃ olefin components are formed. At temperaturesapproaching 450° C., almost complete conversion of methyl bromide tomethane occurs. In the optimum operating temperature range of betweenabout 300° C. and 400° C., as a byproduct of the reaction, a smallamount of carbon will build up on the catalyst over time duringoperation, causing a decline in catalyst activity over a range of hours,up to hundreds of hours, depending on the reaction conditions and thecomposition of the feed gas. It is believed that higher reactiontemperatures above about 400° C., associated with the formation ofmethane favor the thermal cracking of alkyl bromides and formation ofcarbon or coke and hence an increase in the rate of deactivation of thecatalyst. Conversely, temperatures at the lower end of the range,particularly below about 300° C. may also contribute to coking due to areduced rate of desorption of heavier products from the catalyst. Hence,operating temperatures within the range of about 150° C. to about 450°C., but preferably in the range of about 300° C. to about 400° C. in thesecond reactor 134 balance increased selectivity of the desired C₅+products and lower rates of deactivation due to carbon formation,against higher conversion per pass, which minimizes the quantity ofcatalyst, recycle rates and equipment size required.

When a catalyst is selected to form olefins in reactor 134, it ispreferred to operate second reactor 134 at a temperature within therange of about 250° to 500° C. Temperatures above about 450° C. in thesecond reactor can result in increased yields of light hydrocarbons,such as undesirable methane and also deposition of coke, whereas lowertemperatures increase yields of ethylene, propylene, butylene andheavier molecular weight hydrocarbon products. Notably, in the case ofthe alkyl bromide reaction over the preferred 10 X zeolite catalyst, itis believed that cyclization reactions also occur such that the C7+fractions contain substantial substituted aromatics. At increasingtemperatures approaching 400° C., it is believed that methyl bromideconversion increases towards 90% or greater, however selectivity towardsC₅+ products decreases and selectivity towards lighter products,particularly olefins increases. At temperatures exceeding 550° C., it isbelieved that a high conversion of methyl bromide to methane andcarbonaceous, coke occurs. In the preferred operating temperature rangeof between about 300° C. and 450° C., as a byproduct of the reaction, alesser amount of coke probably will build up on the catalyst over timeduring operation, causing a decline in catalyst activity over a range ofhours, up to hundreds of hours, depending on the reaction conditions andthe composition of the feed gas. It is believed that higher reactiontemperatures above about 400° C., associated with the formation ofmethane favor the thermal cracking of alkyl bromides and formation ofcarbon or coke and hence an increase in the rate of deactivation of thecatalyst. Conversely, temperatures at the lower end of the range,particularly below about 300° C. may also contribute to coking due to areduced rate of desorption of heavier products from the catalyst. Hence,operating temperatures within the range of about 250° C. to about 500°C., but preferably in the range of about 300° C. to about 450° C. in thesecond reactor 134 balance increased selectivity of the desired olefinsand C₅+ products and lower rates of deactivation due to carbonformation, against higher conversion per pass, which minimizes thequantity of catalyst, recycle rates and equipment size required.

The catalyst may be periodically regenerated in situ, by isolatingreactor 134 from the normal process flow, purging with an inert gas vialine 170 at a pressure in the range of about 1 bar to about 5 bar and anelevated temperature in the range of 400° C. to 650° C. to removeunreacted material adsorbed on the catalyst insofar as is practical, andthen subsequently oxidizing the deposited carbon to CO₂ by addition ofair or inert gas-diluted oxygen via line 170 to reactor 134 at apressure in the range of about 1 bar to about 5 bar and an elevatedtemperature in the range of 400° C. to 650° C. Carbon dioxide andresidual air or inert gas are vented from reactor 134 via line 175during the regeneration period.

The effluent which comprises hydrobromic acid and higher molecularweight hydrocarbons, olefins or mixtures thereof is withdrawn from thesecond reactor 134 via line 135, cooled to a temperature in the range ofabout 0° C. to about 100° C. in exchanger 136, and combined with vaporeffluent in line 112 from hydrocarbon stripper 147. The mixture is thenpassed to a scrubber 138 and contacted with a stripped, recirculatedwater that is transported to scrubber 138 in line 164 by any suitablemeans, such as pump 143, and is cooled to a temperature in the range ofabout 0° C. to about 50° C. in heat exchanger 155. Any liquidhydrocarbon product condensed in scrubber 138 may be skimmed andwithdrawn as stream 137 and added to liquid hydrocarbon product 154.Hydrobromic acid is dissolved in scrubber 138 in the aqueous solutionwhich is removed from the scrubber 138 via line 144, and passed tohydrocarbon stripper 147 wherein residual hydrocarbons dissolved in theaqueous solution are stripped-out by contact with feed gas 111. Thestripped aqueous phase effluent from hydrocarbon stripper 147 is cooledto a temperature in the range of about 0° C. to about 50° C. in heatexchanger 146 and then passed via line 165 to absorber 148 in whichresidual bromine is recovered from vent stream 167.

The residual vapor phase containing olefins, higher molecular weighthydrocarbons or mixtures thereof is removed as effluent from thescrubber 138 and forwarded to dehydrator 150 to remove substantially allwater from the gas stream. The water is then removed from the dehydrator150 via line 153. The dried gas stream containing olefins, highermolecular weight hydrocarbons or mixtures thereof is further passed vialine 151 to product recovery unit 152 to recover olefins, the C₅+gasoline range hydrocarbon fraction or mixtures thereof as a liquidproduct in line 154. Any conventional method of dehydration and liquidsrecovery such as solid-bed dessicant adsorption followed by, forexample, refrigerated condensation, cryogenic expansion, or circulatingabsorption oil, or other solvents as used to process natural gas orrefinery gas streams and recover olefinic hydrocarbons, as known to askilled artisan, may be employed in the implementation of thisinvention. The residual vapor effluent from product recovery unit 152 isthen split into a purge stream 157 that may be utilized as fuel for theprocess and a recycled residual vapor which is compressed via compressor158. The recycled residual vapor discharged from compressor 158 is splitinto two fractions. A first fraction that is equal to at least 2.5 timesthe feed gas volume is transported via line 162, combined with theliquid bromine conveyed in line 125 and passed to heat exchanger 126wherein the liquid bromine is vaporized and fed into first reactor 130.The second fraction which is drawn off line 162 via line 163 and isregulated by control valve 160, at a rate sufficient to dilute the alkylbromide concentration to reactor 134 and absorb the heat of reactionsuch that reactor 134 is maintained at the selected operatingtemperature, preferably in the range of about 300° C. to about 450° C.in order to maximize conversion vs. selectivity and to minimize the rateof catalyst deactivation due to the deposition of carbon. Thus, thedilution provided by the recycled vapor effluent permits selectivity ofbromination in the first reactor 130 to be controlled in addition tomoderating the temperature in second reactor 134.

Oxygen, oxygen enriched air or air 110 is delivered via blower orcompressor 113 at a pressure in the range of about ambient to about 5bar to bromine stripper 114 to strip residual bromine from water whichleaves stripper 114 via line 164 and is divided into two portions. Thefirst portion of the stripped water is recycled via line 164, cooled inheat exchanger 155 to a temperature in the range of about 20° C. toabout 50° C., and maintained at a pressure sufficient to enter scrubber138 by any suitable means, such as pump 143. The portion of water thatis recycled is selected such that the hydrobromic acid solution effluentremoved from scrubber 138 via line 144 has a concentration in the rangefrom about 10% to about 50% by weight hydrobromic acid, but morepreferably in the range of about 30% to about 48% by weight to minimizethe amount of water which must be vaporized in exchanger 141 andpreheater 119 and to minimize the vapor pressure of HBr over theresulting acid. A second portion of water from stripper 114 is removedfrom line 164 and the process via line 156.

The dissolved hydrobromic acid that is contained in the aqueous solutioneffluent from adsorber 148 is transported via line 149 and is combinedwith the oxygen, oxygen enriched air or air leaving bromine stripper 114in line 115. The combined aqueous solution effluent and oxygen, oxygenenriched air or air is passed to a first side of heat exchanger 141 andthrough preheater 119 wherein the mixture is preheated to a temperaturein the range of about 100° C. to about 600° C. and most preferably inthe range of about 120° C. to about 250° C. and passed to third reactor117 that contains a metal bromide salt or metal oxide. The preferredmetal of the bromide salt or metal oxide is Fe(III), Cu(II) or Zn(II)although Co(II), Ni(II), Mn(II), V(II), Cr(II) or othertransition-metals which form oxidizable bromide salts can be used.Alternatively, alkaline-earth metals which also form oxidizable bromidesalts, such as Ca (II) or Mg(II) could be used. The metal bromide saltin the oxidation reactor 117 can be utilized as a concentrated aqueoussolution or preferably, the concentrated aqueous salt solution may beimbibed into a porous, high surface area, acid resistant inert supportsuch as a silica gel. More preferably, the oxide form of the metal in arange of 10 to 20% by weight is deposited on an inert support such asalumina with a specific surface area in the range of 50 to 200 m2/g. Theoxidation reactor 117 operates at a pressure in the range of aboutambient to about 5 bar and at a temperature in the range of about 100°C. to 600° C., but most preferably in the range of about 130° C. to 350°C.; therein, the metal bromide is oxidized by oxygen, yielding elementalbromine and metal hydroxide, metal oxide or metal oxy-bromide speciesor, metal oxides in the case of the supported metal bromide salt ormetal oxide operated at higher temperatures and lower pressures at whichwater may primarily exist as a vapor. In either case, the hydrobromicacid reacts with the metal hydroxide, metal oxy-bromide or metal oxidespecies and is neutralized, restoring the metal bromide salt andyielding water. Thus, it is believed that the overall reaction resultsin the net oxidation of hydrobromic acid produced in first reactor 130and second reactor 134 to elemental bromine and steam, catalyzed by themetal bromide/metal hydroxide or metal oxide operating in a catalyticcycle. In the case of the metal bromide being Fe(III)Br2 in an aqueoussolution and operated in a pressure and temperature range in which watermay exist as a liquid the reactions are believed to be:Fe(+3a)+6Br(−a)+3H(+a)+3/2O₂(g)=3Br₂(g)+Fe(OH)3  1)3HBr(g)+H₂O=3H(+a)+3Br(−a)+H₂O  2)3H(+a)+3Br(−a)+Fe(OH)3=Fe(+3a)+3Br(−a)+3H₂O  3)In the case of the metal bromide being CU(II)Br2, in an aqueous solutionand operated in a pressure and temperature range in which water mayexist as a liquid the reactions are believed to be:4Cu(+2a)+8Br(−a)+3H₂0+3/2O₂(g)=3Br₂(g)+CuBr₂.3Cu(OH)₂  1)6HBr(g)+H₂O=6H(+a)+6Br(−a)+H₂O  2)6H(+a)+6Br(−a)+CuBr₂.3Cu(OH)₂=4Cu(+2a)+8Br(−a)+6H₂OIn the case of the metal bromide being Cu(II)Br2 supported on an inertsupport and operated at higher temperature and lower pressure conditionsat which water primarily exists as a vapor, the reactions are believedto be:2Cu(II)Br2=2Cu(I)Br+Br2(g)  1)2Cu(I)Br+O₂(g)=Br2(g)+2Cu(II)O  2)2HBr(g)+Cu(II)O=Cu(II)Br₂+H₂O(g)  3)

The elemental bromine and water and any residual oxygen or nitrogen (ifair or oxygen enriched air is utilized as the oxidant) leaving as vaporfrom the outlet of third reactor 117, are cooled in the second side ofexchanger 141 and condenser 120 to a temperature in the range of about0° C. to about 70° C. wherein the bromine and water are condensed andpassed to three-phase separator 122. In three-phase separator 122, sinceliquid water has a limited solubility for bromine, on the order of about3% by weight, any additional bromine which is condensed forms aseparate, denser liquid bromine phase. The liquid bromine phase,however, has a notably lower solubility for water, on the order of lessthan 0.1%. Thus, a substantially dry bromine vapor can be easilyobtained by condensing liquid bromine and water, decanting water bysimple physical separation and subsequently re-vaporizing liquidbromine. It is important to operate at conditions that result in thenear complete reaction of HBr so as to avoid significant residual HBr inthe condensed liquid bromine and water, as HBr increases the miscibilityof bromine in the aqueous phase, and at sufficiently highconcentrations, results in a single ternary liquid phase.

Liquid bromine is pumped from three-phase separator 122 via pump 124 toa pressure sufficient to mix with vapor stream 162. Thus the bromine isrecovered and recycled within the process. The residual air, oxygenenriched air or oxygen and any bromine vapor which is not condensedexits three-phase separator 122 and is passed via line 123 to brominescrubber 148, wherein residual bromine is recovered by dissolution intohydrobromic acid solution stream conveyed to scrubber 148 via line 165.Water is removed from the three-phase separator 122 via line 129 andpassed to stripper 114.

The elemental bromine vapor and steam are condensed and easily separatedin the liquid phase by simple physical separation, yieldingsubstantially dry bromine. The absence of significant water allowsselective bromination of alkanes, without production of CO₂ and thesubsequent efficient and selective reactions of alkyl bromides toprimarily C₂ to C₄ olefins, heavier products, the C₅+ fraction of whichcontains substantial branched alkanes and substituted aromatics, ormixtures thereof. Byproduct hydrobromic acid vapor from the brominationreaction and subsequent reaction in reactor 134 are readily dissolvedinto an aqueous phase and neutralized by the metal hydroxide or metaloxide species resulting from oxidation of the metal bromide.

In accordance with another embodiment of the process of the presentinvention illustrated in FIG. 4A, the alkyl bromination and alkylbromide conversion stages are operated in a substantially similar mannerto those corresponding stages described with respect to FIGS. 2 and 3above. More particularly, a gas stream containing lower molecular weightalkanes, comprised of mixture of a feed gas and a recycled gas stream ata pressure in the range of about 1 bar to about 30 bar, is transportedor conveyed via line, pipe or conduits 262 and 211, respectively, andmixed with dry bromine liquid in line 225. The resultant mixture istransported via pump 224 and passed to heat exchanger 226 wherein theliquid bromine is vaporized. The mixture of lower molecular weightalkanes and dry bromine vapor is fed to reactor 230. Preferably, themolar ratio of lower molecular weight alkanes to dry bromine vapor inthe mixture introduced into reactor 230 is in excess of 2.5:1. Reactor230 has an inlet pre-heater zone 228 which heats the mixture to areaction initiation temperature in the range of 250° C. to 400° C. Infirst reactor 230, the lower molecular weight alkanes are reactedexothermically with dry bromine vapor at a relatively low temperature inthe range of about 250° C. to about 600° C., and at a pressure in therange of about 1 bar to about 30 bar to produce gaseous alkyl bromidesand hydrobromic acid vapors. The upper limit of the operatingtemperature range is greater than the upper limit of the reactioninitiation temperature range to which the feed mixture is heated due tothe exothermic nature of the bromination reaction. In the case ofmethane, the formation of methyl bromide occurs in accordance with thefollowing general reaction:CH₄(g)+Br₂(g)→CH₃Br(g)+HBr(g)This reaction occurs with a significantly high degree of selectivity tomethyl bromide. For example, in the case of bromine reacting with amolar excess of methane at a methane to bromine ratio of 4.5:1, a highselectivity to the mono-halogenated methyl bromide occurs. Small amountsof dibromomethane and tribromomethane are also formed in the brominationreaction. Higher alkanes, such as ethane, propane and butane, are alsoreadily bromoninated resulting in mono and multiple brominated speciessuch as ethyl bromides, propyl bromides and butyl bromides. If an alkaneto bromine ratio of significantly less than 2.5 to 1 is utilized,substantially lower selectivity to methyl bromide occurs and significantformation of undesirable carbon soot is observed. Further, the drybromine vapor that is feed into first reactor 230 is substantiallywater-free. Applicant has discovered that elimination of substantiallyall water vapor from the bromination step in first reactor 230substantially eliminates the formation of unwanted carbon dioxidethereby increasing the selectivity of alkane bromination to alkylbromides and eliminating the large amount of waste heat generated in theformation of carbon dioxide from alkanes.

The effluent that contains alkyl bromides and hydrobromic acid iswithdrawn from the first reactor via line 231 and is partially cooled inheat exchanger 232 before flowing to a second reactor 234. Thetemperature to which the effluent is partially cooled in heat exchanger234 is in the range of about 150° C. to about 350° C. when it is desiredto convert the alkyl bromides to higher molecular weight hydrocarbons insecond reactor 234, or to range of about 150° C. to about 450° C. whenit is desired to convert the alkyl bromides to olefins a second reactor234. In second reactor 234, the alkyl bromides are reactedexothermically over a fixed bed 233 of crystalline alumino-silicatecatalyst, preferably a zeolite catalyst. The temperature and pressureemployed in second reactor, as well as the zeolite catalyst, willdetermine the product that is formed from the reaction of alkyl bromidesoccurring in second reactor 234.

The crystalline alumino-silicate catalyst employed in second reactor 234is preferably a zeolite catalyst and most preferably a ZSM-5 zeolitecatalyst when it is desired to form higher molecular weighthydrocarbons, Although the zeolite catalyst is preferably used in thehydrogen, sodium or magnesium form, the zeolite may also be modified byion exchange with other alkali metal cations, such as Li, Na, K or Cs,with alkali-earth metal cations, such as Mg, Ca, Sr or Ba, or withtransition metal cations, such as Ni, Mn, V, W, or to the hydrogen form.Other zeolite catalysts having varying pore sizes and acidities, whichare synthesized by varying the alumina-to-silica ratio may be used inthe second reactor 234 as will be evident to a skilled artisan.

When it is desired to form olefins from the reaction of alkyl bromidesin reactor 234, the crystalline alumino-silicate catalyst employed insecond reactor 234 is preferably a zeolite catalyst, and most preferablyan X type or Y type zeolite catalyst. A preferred zeolite is 10 X or Ytype zeolite, although other zeolites with differing pore sizes andacidities, which are synthesized by varying the alumina-to-silica ratiomay be used in the process of the present invention as will be evidentto a skilled artisan. Although the zeolite catalyst is preferably usedin a protonic form, a sodium form or a mixed protonic/sodium form, thezeolite may also be modified by ion exchange with other alkali metalcations, such as Li, K or Cs, with alkali-earth metal cations, such asMg, Ca, Sr or Ba, or with transition metal cations, such as Ni, Mn, V,W, or to the hydrogen form. These various alternative cations have aneffect of shifting reaction selectivity. Other zeolite catalysts havingvarying pore sizes and acidities, which are synthesized by varying thealumina-to-silica ratio may be used in the second reactor 234 as will beevident to a skilled artisan.

The temperature at which the second reactor 234 is operated is animportant parameter in determining the selectivity of the reaction tohigher molecular hydrocarbons, or to olefins.

Where a catalyst is selected to form higher molecular weighthydrocarbons in reactor 234, it is preferred to operate second reactor234 at a temperature within the range of about 150° to 450°.Temperatures above about 300° C. in the second reactor result inincreased yields of light hydrocarbons, such as undesirable methane,whereas lower temperatures increase yields of heavier molecular weighthydrocarbon products. At the low end of the temperature range, withmethyl bromide reacting over ZSM-5 zeolite at temperatures as low as150° C. significant methyl bromide conversion on the order of 20% isnoted, with a high selectivity towards C₅+ products. Notably, in thecase of the alkyl bromide reaction over the preferred zeolite ZSM-5catalyst, cyclization reactions also occur such that the C7+ fractionsare composed primarily of substituted aromatics. At increasingtemperatures approaching 300° C., methyl bromide conversion increasestowards 90% or greater, however selectivity towards C₅+ productsdecreases and selectivity towards lighter products, particularlyundesirable methane, increases. Surprisingly, very little ethane orC₂,-C₃ olefin components are formed. At temperatures approaching 450°C., almost complete conversion of methyl bromide to methane occurs. Inthe optimum operating temperature range of between about 300° C. and400° C., as a byproduct of the reaction, a small amount of carbon willbuild up on the catalyst over time during operation, causing a declinein catalyst activity over a range of hours, up to hundreds of hours,depending on the reaction conditions and the composition of the feedgas. It is believed that higher reaction temperatures above about 400°C., associated with the formation of methane favor the thermal crackingof alkyl bromides and formation of carbon or coke and hence an increasein the rate of deactivation of the catalyst. Conversely, temperatures atthe lower end of the range, particularly below about 300° C. may alsocontribute to coking due to a reduced rate of desorption of heavierproducts from the catalyst. Hence, operating temperatures within therange of about 150° C. to about 450° C., but preferably in the range ofabout 300° C. to about 400° C. in the second reactor 234 balanceincreased selectivity of the desired C₅+ products and lower rates ofdeactivation due to carbon formation, against higher conversion perpass, which minimizes the quantity of catalyst, recycle rates andequipment size required.

Where a catalyst is selected to form olefins in reactor 234, it ispreferred to operated second reactor 234 at a temperature within therange of about 250° to 500° C. Temperatures above about 450° C. in thesecond reactor can result in increased yields of light hydrocarbons,such as undesirable methane and also deposition of coke, whereas lowertemperatures increase yields of ethylene, propylene, butylene andheavier molecular weight hydrocarbon products. Notably, in the case ofthe alkyl bromide reaction over the preferred 10 X zeolite catalyst, itis believed that cyclization reactions also occur such that the C7+fractions contain substantial substituted aromatics. At increasingtemperatures approaching 400° C., it is believed that methyl bromideconversion increases towards 90% or greater, however selectivity towardsC₅+ products decreases and selectivity towards lighter products,particularly olefins increases. At temperatures exceeding 550° C., it isbelieved that a high conversion of methyl bromide to methane andcarbonaceous, coke occurs. In the preferred operating temperature rangeof between about 300° C. and 450° C., as a byproduct of the reaction, alesser amount of coke probably will build up on the catalyst over timeduring operation, causing a decline in catalyst activity over a range ofhours, up to hundreds of hours, depending on the reaction conditions andthe composition of the feed gas. It is believed that higher reactiontemperatures above about 400° C., associated with the formation ofmethane favor the thermal cracking of alkyl bromides and formation ofcarbon or coke and hence an increase in the rate of deactivation of thecatalyst. Conversely, temperatures at the lower end of the range,particularly below about 300° C. may also contribute to coking due to areduced rate of desorption of heavier products from the catalyst. Hence,operating temperatures within the range of about 250° C. to about 500°C., but preferably in the range of about 300° C. to about 450° C. in thesecond reactor 234 balance increased selectivity of the desired olefinsand C₅+ products and lower rates of deactivation due to carbonformation, against higher conversion per pass, which minimizes thequantity of catalyst, recycle rates and equipment size required.

The catalyst may be periodically regenerated in situ, by isolatingreactor 234 from the normal process flow, purging with an inert gas vialine 270 at a pressure in the range of about 1 bar to about 5 bar and anelevated temperature in the range of about 400° C. to about 650° C. toremove unreacted material adsorbed on the catalyst insofar as ispractical, and then subsequently oxidizing the deposited carbon to CO₂by addition of air or inert gas-diluted oxygen via line 270 to reactor234 at a pressure in the range of about 1 bar to about 5 bar and anelevated temperature in the range of about 400° C. to about 650° C.Carbon dioxide and residual air or inert gas are vented from reactor 234via line 275 during the regeneration period.

The effluent which comprises hydrobromic acid and higher molecularweight hydrocarbons, olefins or mixtures thereof is withdrawn from thesecond reactor 234 via line 235 and cooled to a temperature in the rangeof about 100° C. to about 600° C. in exchanger 236. As illustrated inFIG. 4A, the cooled effluent is transported via lines 235 and 241 withvalve 238 in the opened position and valves 239 and 243 in the closedposition and introduced into a vessel or reactor 240 containing a bed298 of a solid phase metal oxide. The metal of the metal oxide isselected form magnesium (Mg), calcium (Ca), vanadium (V), chromium (Cr),manganese (Mn), iron (Fe), cobalt (Co), nickel (Ni), copper (Cu), zinc(Sn), or tin (Sn). The metal is selected for the impact of its physicaland thermodynamic properties relative to the desired temperature ofoperation, and also for potential environmental and health impacts andcost. Preferably, magnesium, copper and iron are employed as the metal,with magnesium being the most preferred. These metals have the propertyof not only forming oxides but bromide salts as well, with the reactionsbeing reversible in a temperature range of less than about 500° C. Thesolid metal oxide is preferably immobilized on a suitableattrition-resistant support, for example a synthetic amorphous silica,such as Davicat Grade 57, manufactured by Davison Catalysts of Columbia,Md. Or more preferably, an alumina support with a specific surface areaof about 50 to 200 m2/g. In reactor 240, hydrobromic acid is reactedwith the metal oxide at temperatures below about 600° C. and preferablybetween about 100° C. to about 500° C. in accordance with the followinggeneral formula wherein M represents the metal:2HBr+MO→MBr₂+H₂OThe steam resulting from this reaction is transported together witholefins and/or the high molecular hydrocarbons in line 244, 218 and 216via opened valve 219 to heat exchanger 220 wherein the mixture is cooledto a temperature in the range of about 0° C. to about 70° C. This cooledmixture is forwarded to dehydrator 250 to remove substantially all waterfrom the gas stream. The water is then removed from the dehydrator 250via line 253. The dried gas stream containing olefins, higher molecularweight hydrocarbons or mixtures thereof is further passed via line 251to product recovery unit 252 to recover olefins, the C₅+ fraction, ormixtures thereof as a liquid product in line 254. Any conventionalmethod of dehydration and liquids recovery such as solid-bed dessicantadsorption followed by, for example, refrigerated condensation,cryogenic expansion, or circulating absorption oil or other solvent, asused to process natural gas or refinery gas streams and recover olefinichydrocarbons, as known to a skilled artisan, may be employed in theimplementation of this invention. The residual vapor effluent fromproduct recovery unit 252 is then split into a purge stream 257 that maybe utilized as fuel for the process and a recycled residual vapor whichis compressed via compressor 258. The recycled residual vapor dischargedfrom compressor 258 is split into two fractions. A first fraction thatis equal to at least 1.5 times the feed gas volume is transported vialine 262, combined with the liquid bromine and feed gas conveyed in line225 and passed to heat exchanger 226 wherein the liquid bromine isvaporized and fed into first reactor 230 in a manner as described above.The second fraction which is drawn off line 262 via line 263 and isregulated by control valve 260, at a rate sufficient to dilute the alkylbromide concentration to reactor 234 and absorb the heat of reactionsuch that reactor 234 is maintained at the selected operatingtemperature, preferably in the range of about 300° C. to about 450° C.in order to maximize conversion vs. selectivity and to minimize the rateof catalyst deactivation due to the deposition of carbon. Thus, thedilution provided by the recycled vapor effluent permits selectivity ofbromination in the first reactor 230 to be controlled in addition tomoderating the temperature in second reactor 234.

Oxygen, oxygen enriched air or air 210 is delivered via blower orcompressor 213 at a pressure in the range of about ambient to about 10bar to bromine via line 214, line 215 and valve 249 through heatexchanger 215, wherein oxygen, oxygen enriched air or air is preheatedto a temperature in the range of about 100° C. to about 500° C. to asecond vessel or reactor 246 containing a bed 299 of a solid phase metalbromide. Oxygen reacts with the metal bromide in accordance with thefollowing general reaction wherein M represents the metal:MBr₂+½O₂→MO+Br₂In this manner, a dry, substantially HBr free bromine vapor is producedthereby eliminating the need for subsequent separation of water orhydrobromic acid from the liquid bromine. Reactor 246 is operated below600° C., and more preferably between about 300° C. to about 500° C. Theresultant bromine vapor is transported from reactor 246 via line 247,valve 248 and line 242 to heat exchanger or condenser 221 where thebromine is condensed into a liquid. The liquid bromine is transportedvia line 242 to separator 222 wherein liquid bromine is removed via line225 and transported via line 225 to heat exchanger 226 and first reactor230 by any suitable means, such as by pump 224. The residual air orunreacted oxygen is transported from separator 222 via line 227 to abromine scrubbing unit 223, such as venturi scrubbing system containinga suitable solvent, or suitable solid adsorbant medium, as selected by askilled artisan, wherein the remaining bromine is captured. The capturedbromine is desorbed from the scrubbing solvent or adsorbant by heatingor other suitable means and the thus recovered bromine transported vialine 212 to line 225. The scrubbed air or oxygen is vented via line 229.In this manner, nitrogen and any other substantially non-reactivecomponents are removed from the system of the present invention andthereby not permitted to enter the hydrocarbon-containing portion of theprocess; also loss of bromine to the surrounding environment is avoided.

One advantage of removing the HBr by chemical reaction in accordancewith this embodiment, rather than by simple physical solubility, is thesubstantially complete scavenging of the HBr to low levels at higherprocess temperatures. Another distinct advantage is the elimination ofwater from the bromine removed thereby eliminating the need forseparation of bromine and water phases and for stripping of residualbromine from the water phase.

Reactors 240 and 246 may be operated in a cyclic fashion. As illustratedin FIG. 4A, valves 238 and 219 are operated in the open mode to permithydrobromic acid to be removed from the effluent that is withdrawn fromthe second reactor 234, while valves 248 and 249 are operated in theopen mode to permit air, oxygen enriched air or oxygen to flow throughreactor 246 to oxidize the solid metal bromide contained therein. Oncesignificant conversion of the metal oxide and metal bromide in reactors240 and 246, respectively, has occurred, these valves are closed. Atthis point, bed 299 in reactor 246 is a bed of substantially solid metalbromide, while bed 298 in reactor 240 is substantially solid metaloxide. As illustrated in FIG. 5A, valves 245 and 243 are then opened topermit oxygen, oxygen enriched air or air to flow through reactor 240 tooxidize the solid metal bromide contained therein, while valves 239 and217 are opened to permit effluent which comprises olefins, the highermolecular weight hydrocarbons and/or hydrobromic acid that is withdrawnfrom the second reactor 234 to be introduced into reactor 246. Thereactors are operated in this manner until significant conversion of themetal oxide and metal bromide in reactors 246 and 240, respectively, hasoccurred and then the reactors are cycled back to the flow schematicillustrated in FIG. 4A by opening and closing valves as previouslydiscussed.

When oxygen is utilized as the oxidizing gas transported in via line 210to the reactor being used to oxidize the solid metal bromide containedtherein, the embodiment of the process of the present inventionillustrated in FIGS. 4A and 5A can be modified such that the brominevapor produced from either reactor 246 (FIG. 4B) or 240 (FIG. 5B) istransported via lines 242 and 225 directly to first reactor 230. Sinceoxygen is reactive and will not build up in the system, the need tocondense the bromine vapor to a liquid to remove unreactive components,such as nitrogen, is obviated. Compressor 213 is not illustrated inFIGS. 4B and 5B as substantially all commercial sources of oxygen, suchas a commercial air separator unit, will provide oxygen to line 210 atthe required pressure. If not, a compressor 213 could be utilized toachieve such pressure as will be evident to a skilled artisan.

In the embodiment of the present invention illustrated in FIG. 6A, thebeds of solid metal oxide particles and solid metal bromide particlescontained in reactors 240 and 246, respectively, are fluidized and areconnected in the manner described below to provide for continuousoperation of the beds without the need to provide for equipment, such asvalves, to change flow direction to and from each reactor. In accordancewith this embodiment, the effluent which comprises olefins, the highermolecular weight hydrocarbons and/or hydrobromic acid is withdrawn fromthe second reactor 234 via line 235, cooled to a temperature in therange of about 100° C. to about 500° C. in exchanger 236, and introducedinto the bottom of reactor 240 which contains a bed 298 of solid metaloxide particles. The flow of this introduced fluid should induce theparticles in bed 298 to move upwardly within reactor 240 as thehydrobromic acid is reacted with the metal oxide in the manner asdescribed above with respect to FIG. 4A. At or near the top of the bed298, the particles which contain substantially solid metal bromide onthe attrition-resistant support due to the substantially completereaction of the solid metal oxide with hydrobromic acid in reactor 240are withdrawn via a weir or cyclone or other conventional means ofsolid/gas separation, flow by gravity down line 259 and are introducedat or near the bottom of the bed 299 of solid metal bromide particles inreactor 246. In the embodiment illustrated in FIG. 6A, oxygen, oxygenenriched air or air 210 is delivered via blower or compressor 213 at apressure in the range of about ambient to about 10 bar, transported vialine 214 through heat exchanger 215, wherein the oxygen, oxygen enrichedair or air is preheated to a temperature in the range of about 100° C.to about 500° C. and introduced into second vessel or reactor 246 belowbed 299 of a solid phase metal bromide. Oxygen reacts with the metalbromide in the manner described above with respect to FIG. 4A to producea dry, substantially HBr free bromine vapor. The flow of this introducedgas should induce the particles in bed 299 to flow upwardly withinreactor 246 as oxygen is reacted with the metal bromide. At or near thetop of the bed 298, the particles which contain substantially solidmetal oxide on the attrition-resistant support due to the substantiallycomplete reaction of the solid metal bromide with oxygen in reactor 246are withdrawn via a weir or cyclone or other conventional means ofsolid/gas separation, flow by gravity down line 264 and are introducedat or near the bottom of the bed 298 of solid metal oxide particles inreactor 240. In this manner, reactors 240 and 246 may be operatedcontinuously without changing the parameters of operation.

In the embodiment illustrated in FIG. 6B, oxygen is utilized as theoxidizing gas and is transported in via line 210 to reactor 246.Accordingly, the embodiment of the process of the present inventionillustrated in FIG. 6A is modified such that the bromine vapor producedfrom reactor 246 is transported via lines 242 and 225 directly to firstreactor 230. Since oxygen is reactive and will not build up in thesystem, it is believed that the need to condense the bromine vapor to aliquid to remove unreactive components, such as nitrogen, should beobviated. Compressor 213 is not illustrated in FIG. 6B as substantiallyall commercial sources of oxygen, such as a commercial air separatorunit, will provide oxygen to line 210 at the required pressure. If not,a compressor 213 could be utilized to achieve such pressure as will beevident to a skilled artisan.

In accordance with another embodiment of the process of the presentinvention that is illustrated in FIG. 7, the alkyl bromination and alkylbromide conversion stages are operated in a substantially similar mannerto those corresponding stages described in detail with respect to FIG.4A except as discussed below. Residual air or oxygen and bromine vaporemanating from reactor 246 is transported via line 247, valve 248 andline 242 and valve 300 to heat exchanger or condenser 221 wherein thebromine-containing gas is cooled to a temperature in the range of about30° C. to about 300° C. The bromine-containing vapor is then transportedvia line 242 to vessel or reactor 320 containing a bed 322 of a solidphase metal bromide in a reduced valence state. The metal of the metalbromide in a reduced valence state is selected from copper (Cu), iron(Fe), or molybdenum (Mo). The metal is selected for the impact of itsphysical and thermodynamic properties relative to the desiredtemperature of operation, and also for potential environmental andhealth impacts and cost. Preferably, copper or iron are employed as themetal, with copper being the most preferred. The solid metal bromide ispreferably immobilized on a suitable attrition-resistant support, forexample a synthetic amorphous silica, such as Davicat Grade 57,manufactured by Davison Catalysts of Columbia, Md. More preferably themetal is deposited in oxide form in a range of about 10 to 20 wt % on analumina support with a specific surface area in the range of about 50 to200 m2/g, In reactor 320, bromine vapor is reacted with the solid phasemetal bromide, preferably retained on a suitable attrition-resistantsupport at temperatures below about 300° C. and preferably between about30° C. to about 200° C. in accordance with the following general formulawherein M² represents the metal:2M²Br_(n)+Br₂→2M²Br_(n+1)In this manner, bromine is stored as a second metal bromide, i.e.2M²Br_(n+1), in reactor 320 while the resultant vapor containingresidual air or oxygen is vented from reactor 320 via line 324, valve326 and line 318.

The gas stream containing lower molecular weight alkanes, comprised ofmixture of a feed gas (line 211) and a recycled gas stream, istransported or conveyed via line 262, heat exchanger 352, wherein thegas stream is preheated to a temperature in the range of about 150° C.to about 600° C., valve 304 and line 302 to a second vessel or reactor310 containing a bed 312 of a solid phase metal bromide in an oxidizedvalence state. The metal of the metal bromide in an oxidized valencestate is selected from copper (Cu), iron (Fe), or molybdenum (Mo). Themetal is selected for the impact of its physical and thermodynamicproperties relative to the desired temperature of operation, and alsofor potential environmental and health impacts and cost. Preferably,copper or iron are employed as the metal, with copper being the mostpreferred. The solid metal bromide in an oxidized state is preferablyimmobilized on a suitable attrition-resistant support, for example asynthetic amorphous silica such as Davicat Grade 57, manufactured byDavison Catalysts of Columbia, Md. More preferably the metal isdeposited in an oxide state in a range of 10 to 20 wt % supported on analumina support with a specific surface area of about 50 to 200 m2/g.The temperature of the gas stream is from about 150° C. to about 600°C., and preferably from about 200° C. to about 450° C. In second reactor310, the temperature of the gas stream thermally decomposes the solidphase metal bromide in an oxidized valence state to yield elementalbromine vapor and a solid metal bromide in a reduced state in accordancewith the following general formula wherein M² represents the metal:2M²Br_(n+1)→2M²Br_(n)+Br₂The resultant bromine vapor is transported with the gas streamcontaining lower molecular weight alkanes via lines 314, 315, valve 317,line 330, heat exchanger 226 prior to being introduced into alkylbromination reactor 230.

Reactors 310 and 320 may be operated in a cyclic fashion. As illustratedin FIG. 7, valve 304 is operated in the open mode to permit the gasstream containing lower molecular weight alkanes to be transported tothe second reactor 310, while valve 317 is operated in the open mode topermit this gas stream with bromine vapor that is generated in reactor310 to be transported to alkyl bromination reactor 230. Likewise, valve306 is operated in the open mode to permit bromine vapor from reactor246 to be transported to reactor 320, while valve 326 is operated in theopen mode to permit residual air or oxygen to be vented from reactor320. Once significant conversion of the reduced metal bromide andoxidized metal bromide in reactors 320 and 310, respectively, to thecorresponding oxidized and reduced states has occurred, these valves areclosed as illustrated in FIG. 8. At this point, bed 322 in reactor 320is a bed of substantially metal bromide in an oxidized state, while bed312 in reactor 310 is substantially metal bromide in a reduced state. Asillustrated in FIG. 8, valves 304, 317, 306 and 326 are closed, and thenvalves 308 and 332 are opened to permit the gas stream containing lowermolecular weight alkanes to be transported or conveyed via lines 262,heat exchanger 352, wherein gas stream is heated to a range of about150° C. to about 600° C., valve 308 and line, 309 to reactor 320 tothermally decompose the solid phase metal bromide in an oxidized valencestate to yield elemental bromine vapor and a solid metal bromide in areduced state. Valve 332 is also opened to permit the resultant brominevapor to be transported with the gas stream containing lower molecularweight alkanes via lines 324 and 330 and heat exchanger 226 prior tobeing introduced into alkyl bromination reactor 230. In addition, valve300 is opened to permit. bromine vapor emanating from reactor 246 to betransported via line 242 through exchanger 221 into reactor 310 whereinthe solid phase metal bromide in a reduced valence state reacts withbromine to effectively store bromine as a metal bromide. In addition,valve 316 is opened to permit the resulting gas, which is substantiallydevoid of bromine to be vented via lines 314 and 318. The reactors areoperated in this manner until significant conversion of the beds ofreduced metal bromide and oxidized metal bromide in reactors 310 and320, respectively, to the corresponding oxidized and reduced states hasoccurred and then the reactors are cycled back to the flow schematicillustrated in FIG. 7 by opening and closing valves as previouslydiscussed.

In the embodiment of the present invention illustrated in FIG. 9, thebeds 312 and 322 contained in reactors 310 and 320, respectively, arefluidized and are connected in the manner described below to provide forcontinuous operation of the beds without the need to provide forequipment, such as valves, to change flow direction to and from eachreactor. In accordance with this embodiment, the bromine-containing gaswithdrawn from the reactor 246 via line 242 is cooled to a temperaturein the range of about 30° C. to about 300° C. in exchangers 370 and 372,and introduced into the bottom of reactor 320 which contains a movingsolid bed 322 in a fluidized state. The flow of this introduced fluidshould induce the particles in bed 322 to flow upwardly within reactor320 as the bromine vapor is reacted with the reduced metal bromideentering the bottom of bed 322 in the manner as described above withrespect to FIG. 7. At or near the top of the bed 322, the particleswhich contain substantially oxidized metal bromide on theattrition-resistant support due to the substantially complete reactionof the reduced metal bromide with bromine vapor in reactor 320 arewithdrawn via a weir, cyclone or other conventional means of solid/gasseparation, flow by gravity down line 359 and are introduced at or nearthe bottom of the bed 312 in reactor 310. In the embodiment illustratedin FIG. 9, the gas stream containing lower molecular weight alkanes,comprised of mixture of a feed gas (line 211) and a recycled gas stream,is transported or conveyed via line 262 and heat exchanger 352 whereinthe gas stream is heated to a range of about 150° C. to about 600° C.and introduced into reactor 310. The heated gas stream thermallydecomposes the solid phase metal bromide in an oxidized valence statepresent entering at or near the bottom of bed 312 to yield elementalbromine vapor and a solid metal bromide in a reduced state. The flow ofthis introduced gas should induce the particles in bed 312 to flowupwardly within reactor 310 as the oxidized metal bromide is thermallydecomposed. At or near the top of the bed 312, the particles whichcontain substantially reduced solid metal bromide on theattrition-resistant support due to the substantially complete thermaldecomposition in reactor 310 are withdrawn via a weir or cyclone orother conventional means of gas/solid separation and flow by gravitydown line 364 and introduced at or near the bottom of the bed 322 ofparticles in reactor 310. In this manner, reactors 310 and 320 may beoperated continuously with changing the parameters of operation.

It is believed that the process of the present invention should be lessexpensive than conventional process since it operates at low pressuresin the range of about 1 bar to about 30 bar and at relatively lowtemperatures in the range of about 20° C. to about 600° C. for the gasphase, and preferably about 20° C. to about 180° C. for the liquidphase. It is believed that these operating conditions should permit theuse of less expensive equipment of relatively simple design that areconstructed from readily available metal alloys or glass-lined equipmentfor the gas phase and polymer-lined or glass-lined vessels, piping andpumps for the liquid phase. It is believed that the process of thepresent invention also should be more efficient because less energyshould be required for operation and the production of excessive carbondioxide as an unwanted byproduct is minimized. The process is capable ofdirectly producing a mixed hydrocarbon product containing variousmolecular-weight components in the liquefied petroleum gas (LPG), olefinand motor gasoline fuels range that have substantial aromatic contentthereby significantly increasing the octane value of the gasoline-rangefuel components.

The following examples demonstrate the practice and utility of thepresent invention, but are not to be construed as limiting the scopethereof.

EXAMPLE 1

Various mixtures of dry bromine and methane are reacted homogeneously attemperatures in the range of 459° C. to 491° C. at a Gas Hourly SpaceVelocity (GHSV which is defined as the gas flow rate in standard litersper hour divided by the gross reactor catalyst-bed volume, includingcatalyst-bed porosity, in liters) of approximately 7200 hr⁻¹. Theresults of this example indicate that for molar ratios of methane tobromine greater than 4.5:1 selectivity to methyl bromide is in the rangeof 90 to 95%, with near-complete conversion of bromine.

EXAMPLE 2

FIG. 13 and FIG. 14 illustrate two exemplary PONA analyses of two C₆+liquid product samples that are recovered during two test runs withmethyl bromide and methane reacting over ZSM-5 zeolite catalyst. Theseanalyses show the substantially aromatic content of the C₆+ fractionsproduced.

EXAMPLE 3

Methyl bromide is reacted over a ZSM-5 zeolite catalyst at a Gas HourlySpace Velocity (GHSV) of approximately 94 hr⁻¹ over a range oftemperatures from about 100° C. to about 460° C. at approximately 2 barpressure. As illustrated in FIG. 10, which is a graph of methyl bromideconversion and product selectivity for the oligimerization reaction as afunction of temperature, methyl bromide conversion increases rapidly inthe range of about 200° C. to about 350° C. Lower temperatures in therange of about 100° C. to about 250° C. favor selectivity towards highermolecular weight products however conversion is low. Higher temperaturesin the range of about 250° C. to about 350° C. show higher conversionsin the range of 50% to near 100%, however increasing selectivity tolower molecular weight products, in particular undesirable methane isobserved. At higher temperatures above 350° C. selectivity to methanerapidly increases. At about 450° C., almost complete conversion tomethane occurs.

EXAMPLE 4

Methyl bromide, hydrogen bromide and methane are reacted over a ZSM-5zeolite catalyst at approximately 2 bar pressure at about 250° C. andalso at about 260° C. at a GHSV of approximately 76 hr⁻¹. Comparisontests utilizing a mixture of only methyl bromide and methane withouthydrogen bromide over the same ZSM-5 catalyst at approximately the samepressure at about 250° C. and at about 260° C. at a GHSV ofapproximately 73 hr⁻¹ were also run. FIG. 11, which is a graph thatillustrates the comparative conversions and selectivities of severalexample test runs, shows only a very minor effect due to the presence ofHBr on product selectivities. Because hydrobromic acid has only a minoreffect on conversion and selectivity, it is not necessary to remove thehydrobromic acid generated in the bromination reaction step prior to theconversion reaction of the alkyl bromides, in which additionalhydrobromic acid is formed in any case. Thus, the process can besubstantially simplified.

EXAMPLE 5

Methyl bromide is reacted over a ZSM-5 zeolite catalyst at 230° C.Dibromomethane is added to the reactor. FIG. 12, which is a graph ofproduct selectivity, indicates that reaction of methyl bromide anddibromomethane results in a shift in selectivity towards C₅+ productsversus. methyl bromide alone. Thus, these results demonstrate thatdibromomethane is also reactive and therefore very high selectivity tobromomethane in the bromination step is not required in the process ofthe present invention. It has been observed, however, that the presenceof dibromomethane increases the rate of catalyst deactivation, requiringa higher operating temperature to optimize the tradeoff betweenselectivity and deactivation rate, as compared to pure methyl bromide.

EXAMPLE 6

A mixture of 12.1 mol % methyl bromide and 2.8 mol % propyl bromide inmethane are reacted over a ZSM-5 zeolite catalyst at 295 C and a GHSV ofapproximately 260 hr⁻¹. A methyl bromide conversion of approximately 86%and a propyl bromide conversion of approximately 98% is observed.

Thus, in accordance with all embodiments of the present invention setforth above, the metal bromide/metal hydroxide, metal oxy-bromide ormetal oxide operates in a catalytic cycle allowing bromine to be easilyrecycled within the process. The metal bromide is readily oxidized byoxygen, oxygen enriched air or air either in the aqueous phase or thevapor phase at temperatures in the range of about 100° C. to about 600°C. and most preferably in the range of about 120° C. to about 180° C. toyield elemental bromine vapor and metal hydroxide, metal oxy-bromide ormetal oxide. Operation at temperatures below about 180° C. isadvantageous, thereby allowing the use of low-cost corrosion-resistantfluoropolymer-lined equipment. Hydrobromic acid is neutralized byreaction with the metal hydroxide or metal oxide yielding steam and themetal bromide.

While the foregoing preferred embodiments of the invention have beendescribed and shown, it is understood that the alternatives andmodifications, such as those suggested and others, may be made theretoand fall within the scope of the invention.

I claim:
 1. A process comprising: separating hydrobromic acid from agaseous stream comprising hydrobromic acid and hydrocarbons; convertingsaid hydrobromic acid to at least bromine; contacting said bromine withgaseous alkanes, in a first reactor, to form bromination productscomprising alkyl bromides; reacting said alkyl bromides in a secondreactor to convert at least a portion of said alkyl bromides tosynthesis products comprising hydrocarbon products and additionalhydrobromic acid, wherein the hydrocarbon products comprise hydrocarbonsselected from the group consisting of olefins, C5+ hydrocarbons, andcombinations thereof; withdrawing effluent from the second reactorcomprising the hydrocarbon products and additional hydrobromic acid; andconverting the additional hydrobromic acid in the effluent withdrawnfrom the second reactor to at least bromine.
 2. The process of claim 1wherein said step of separating said hydrobromic acid from saidhydrocarbons comprises: contacting said gaseous stream with water. 3.The process of claim 2 wherein said step of contacting said gaseousstream with said water comprises: neutralizing said hydrobromic acid toform an aqueous solution comprising said water and a metal bromide salt,the metal of said metal bromide salt being selected from Cu, Zn, Fe, Co,Ni, Mn, Ca or Mg bromide.
 4. The process of claim 3 wherein said step ofconverting comprises: oxidizing said aqueous solution containing saidmetal bromide salt to form at least said bromine and a reaction productselected from the group consisting of a metal hydroxide, a metaloxy-bromide species and combinations thereof; and separating saidbromine from said reaction product.
 5. The process of claim 4 whereinsaid water that contacts said gaseous stream comprises said reactionproduct.
 6. The process of claim 2 wherein said hydrobromic aciddissolves into said water forming a hydrobromic acid solution, saidprocess further comprising: vaporizing said hydrobromic acid solution;and reacting said vaporized hydrobromic acid solution with a metal oxideto form a reaction product comprising a metal bromide salt, the metal ofsaid metal bromide salt being selected from the group of Cu, Zn, Fe, Co,Ni, Mn, Ca or Mg.
 7. The process of claim 6 wherein said step ofconverting comprises: oxidizing said metal bromide salt to formoxidation products comprising said bromine and said metal oxide; andseparating said bromine from said metal oxide.
 8. The process of claim 6wherein said metal bromide salt is contained on a porous support.
 9. Theprocess of claim 1 wherein said step of separating hydrobromic acid fromhydrocarbons comprises reacting said hydrobromic acid with a metal oxideto form reaction products comprising a metal bromide and steam.
 10. Theprocess of claim 9 wherein the metal of said metal oxide is magnesium,calcium, vanadium, chromium, manganese, iron, cobalt, nickel, copper,zinc or tin.
 11. The process of claim 10 wherein said metal oxide issupported on a solid carrier.
 12. The process of claim 11 wherein saidmetal oxide is contained in a bed in a vessel.
 13. The process of claim9 wherein said step of converting comprises: reacting said metal bromidewith an oxygen containing gas to obtain reaction products comprisingsaid metal oxide and said bromine.
 14. The process of claim 1 whereinsaid bromine from the reaction of said metal bromide with said oxygencontaining gas is recycled to said step of contacting said gaseousalkanes to form alkyl bromides.
 15. The process of claim 1 wherein thestep of reacting said alkyl bromides is in the presence a syntheticcrystalline alumino-silicate catalyst.
 16. A process comprising:contacting a gaseous stream comprising hydrobromic acid and hydrocarbonswith an aqueous solution comprising a base selected from the groupconsisting of a metal hydroxide, a metal oxy-bromide species, andcombinations thereof such that the hydrobromic acid is neutralized toform a metal bromide salt in the aqueous solution; oxidizing saidaqueous solution containing said metal bromide salt to form oxidationproducts comprising bromine and said base; separating said bromine fromsaid aqueous solution comprising said base; and contacting said brominewith gaseous alkanes to form alkyl bromides.
 17. The process of claim 16further comprising: reacting said alkyl bromides in the presence of saidhydrobromic acid and a synthetic crystalline alumino-silicate catalystto form said hydrocarbons.
 18. A process comprising: contacting agaseous stream comprising hydrobromic acid and hydrocarbons with water,wherein said hydrobromic acid dissolves in said water to form an aqueoussolution comprising said water and said hydrobromic acid; neutralizingsaid hydrobromic acid to form a metal bromide salt; oxidizing said metalbromide salt to form an oxidation product comprising bromine; andcontacting said bromine with gaseous alkanes to form brominationproducts comprising alkyl bromides.
 19. The process of claim 18 furthercomprising: reacting said alkyl bromides in the presence of saidhydrobromic acid and a synthetic crystalline alumino-silicate catalystto form reaction products comprising said hydrocarbons.
 20. A processcomprising: reacting hydrobromic acid with a metal oxide to formreaction products comprising a metal bromide and steam, wherein saidhydrobromic acid is contained in a gaseous stream comprising saidhydrobromic acid and hydrocarbons; reacting said metal bromide with agas comprising oxygen to form reaction products comprising bromine andsaid metal oxide; contacting said bromine with gaseous alkanes, in afirst reactor, to form bromination products comprising alkyl bromides,reacting said alkyl bromides in a second reactor to convert at least aportion of said alkyl bromides to synthesis products comprisinghydrocarbon products and additional hydrobromic acid, wherein thehydrocarbon products comprise hydrocarbons selected from the groupconsisting of olefins, C5+ hydrocarbons, and combinations thereof;withdrawing effluent from the second reactor comprising the hydrocarbonproducts and additional hydrobromic acid; and reacting the additionalhydrobromic acid in the effluent withdrawn from the second reactor witha metal oxide.
 21. The process of claim 19 wherein the step of reactingsaid alkyl bromides is in the presence of a synthetic crystallinealumino-silicate catalyst.